Method for the hydrogenation of maleic anhydride and related compounds in two serial reaction zones

ABSTRACT

The present invention relates to a process for the gas-phase hydrogenation of C 4 -dicarboxylic acids and/or their derivatives over a catalyst based on copper oxide to give substituted or unsubstituted γ-butyrolactone and/or tetrahydrofuran, which comprises, which a first reaction zone in which the C 4 -dicarboxylic acid and/or its derivatives is/are reacted to give a mixture comprising a substituted or unsubstituted γ-butyrolactone as main product and a subsequent second reaction zone in which the substituted or unsubstituted γ-butyrolactone present in the mixture from the first hydrogenation step is reacted at a temperature lower than the temperature in the first hydrogenation step to give substituted or unsubstituted tetrahydrofuran.

The present invention relates to a process for preparing unsubstitutedor alkyl-substituted γ-butyrolactone and tetrahydrofuran by catalytichydrogenation in the gas phase of substrates selected from the groupconsisting of maleic acid and succinic acid and derivatives of theseacids. For the purposes of the present invention, such derivatives areesters and anhydrides which, like the acids themselves, may bear one ormore alkyl substituents. The process makes it possible to achieve highyields and the ratio of the two products can be set over a wide range.The process of the present invention is carried out in two reactionzones connected in series.

The preparation of γ-butyrolactone (GBL) and tetrahydrofuran THF) bygas-phase hydrogenation of maleic anhydride (MA) is a reaction which hasbeen known for many years. Numerous catalyst systems for carrying outthis catalytic reaction are described in the literature. These aremostly Cr-containing. Depending on the composition of the catalyst andthe reaction parameters selected, such catalysts give different productdistributions.

Apart from MA, further possible starting materials for preparing GBL andTHF are maleic acid itself, succinic acid and its anhydride and alsoesters of these acids. If GBL and THF bearing alkyl substituents are tobe prepared, the appropriately alkyl-substituted derivatives of theabovementioned acids, esters and anhydrides can be used.

U.S. Pat. No. 3,065,243 discloses a process in which copper chromite isused as catalyst. According to the description and examples, thisprocess forms considerable amounts of succinic anhydride (SA) which hasto be circulated. As is well known, this frequently results in processengineering problems due to crystallization of SA or succinic acidformed therefrom followed by blockage of pipes.

Further copper chromite catalysts for the hydrogenation of MA aredisclosed, for example, in the publications U.S. Pat. Nos. 3,580,930,4,006,165, EP-A 638 565 and WO 99/38856. According to these disclosures,high yields of GBL can be achieved using the catalysts described there.THF is in each case formed only in traces. However, larger amounts ofTHF are often desired for a number of reasons.

A process which allows this is disclosed in U.S. Pat. No. 5,072,009. Thecatalysts used according to this patent have the formulaCu₁Zn_(b)Al_(c)M_(d)O_(x), where M is at least one element selected fromthe group consisting of the elements of groups IIA and IIIA, VA, VIII,Ag, Au, groups IIIB to VIIB and the lanthanides and actinides of thePeriodic Table of the Elements, b is from 0.001 to 500, c is from 0.001to 500 and d is from 0 to <200 and x corresponds to the number of oxygenatoms required according to the valence criteria. Although it is statedthat it is not necessary for chromium to be present in the catalystsdisclosed in this patent, all examples describe chromium-containingcatalysts. According to these examples, the maximum THF yield is 96%,and the hydrogenation is carried out at pressures of from 20 to 40 bar.

An in-principle disadvantage of all the above-described catalyst systemsis the presence of chromium oxide whose use should be avoided because ofthe high toxicity. Cr-free catalyst systems for preparing GBL byhydrogenation of MA have also been described in the prior art. Examplesof such catalyst systems may be found in the publications WO 99/35139(Cu—Zn oxide), WO 95/22539 (Cu—Zn—Zr) and U.S. Pat. No. 5,122,495(Cu—Zn—Al oxide). All these catalyst systems make it possible to achievehigh yields of GBL, up to 98%, but the formation of THF is not observedor only traces are formed. Although the formation of THF can, as isknown, be promoted by increasing the reaction temperature or having alonger residence time in the reactor, this at the same time alsoincreases the proportion of undesirable by-products, for examplebutanol, butane, ethanol or ethane.

A catalyst made up exclusively of copper and aluminum oxides for thegas-phase hydrogenation of MA to GBL is disclosed in WO 97/24346. Heretoo, the same disadvantages as in the publications described in theprevious paragraph, namely formation of THF in only minor amounts ortraces, are encountered.

The use of a catalyst having essentially the same composition asdescribed in WO 97/24346, namely based on Cu—Al oxides, is alsodisclosed in JP 2 233 631. The object of that invention is to carry outthe hydrogenation of MA in such a way that THF and 1,4-butanediol areformed as main products and only small amounts, if any, of GBL areformed. This is achieved by the use of the catalysts based on mixedCu—Al oxides and by adhering to particular reaction conditions. Typicalmixtures obtained by means of this process comprise from about 15 to 20mol % of 1,4-butanediol and from 60 to 80 mol % of THF, with the amountof THF even being able to be increased to over 99 mol % according to oneexample. This is achieved by using GBL as solvent in a large excess. If,on the other hand, no solvent is employed, the yields drop significantlyto values of 75%.

In contrast, EP-A 0 404 408 discloses a catalyst for the hydrogenationof MA whose structure is different in principle from that of thecatalysts in the above-mentioned references. Here, the catalyticallyactive material corresponds essentially to the material disclosed in theabove-cited U.S. Pat. No. 5,072,009. The material is then applied to anessentially inert, at least partly porous support having an externalsurface. The catalytically active material adheres to the outer surfaceof the support. In contrast to the corresponding, unsupported catalyst,which gives THF as main product, this catalyst forms GBL as preferredproduct. Here too, Cr is present in all catalysts used in the examples.Another disadvantage is the large amount of SA formed.

All the types of catalyst described in the abovementioned publicationshave the disadvantage that they still produce a large amount ofundesired by-product or can be used only for the preparation of one ofthe main products THF and GBL which may be desired in principle. Inaddition, Cr is frequently present in the catalysts.

A two-stage process for the hydrogenation of MA is described in U.S.Pat. No. 5,149,836. This process enables GBL and THF to be produced inan adjustable selectivity ratio of from 15 to 92% of GBL or from 7 to83% of THF. The process comprises a first step in which MA ishydrogenated over a first catalyst bed comprising from 30 to 65% byweight of CuO, from 18 to 50% by weight of ZnO and from 8 to 22% byweight of Al₂O₃ to give a gas mixture comprising predominantly GBL. TheGBL obtained in the first step is hydrogenated to THF over a secondcatalyst bed comprising from 10 to 50% by weight of CuO, from 30 to 65%by weight of ZnO and from 3 to 20% by weight of Cr₂O₃. The firsthydrogenation is carried out at from 200 to 400° C., while the second iscarried out at from 200 to 350° C., preferably from 250 to 280° C.According to the examples, the reaction temperature in the first step isfrom 245 to 275° C. while that in the second step is from 250 to 280° C.At a temperature of 250° C., mainly butanediol is formed in the secondhydrogenation step, while mainly THF is formed at 280° C.

WO 99/35136 describes a further process for preparing THF and GBL invariable relative amounts. Starting materials used are maleic anhydrideor succinic anhydride or fumaric esters. In a first step, these arereacted with hydrogen over a copper-based heterogeneous catalyst,preferably a copper-zinc oxide or stabilized copper chromite catalyst.In a second reaction step, an acidic silicon-aluminum oxide is used.Disadvantages of this process are the use of two entirely differentcatalysts and also the limited flexibility in respect of the productmix, since the GBL:THF ratio can be varied only in the range from 70:30to 40:60.

It is an object of the present invention to provide a process for thegas-phase hydrogenation of maleic acid and/or succinic acid and/or theabovementioned derivatives by means of which substituted orunsubstituted GBL and/or THF can be prepared and which allows these twoproducts to be prepared in widely variable relative amounts and in highyields.

We have found that this object is achieved by a process for thegas-phase hydrogenation of C₄-dicarboxylic acids and/or theirderivatives over a catalyst based on copper oxide to give substituted orunsubstituted γ-butyrolactone and/or tetrahydrofuran, which processcomprises a first hydrogenation step in which the C₄-dicarboxylic acidor its derivative is reacted to give a mixture comprising a substitutedor unsubstituted γ-butyrolactone as main product and a subsequent secondhydrogenation step in which the substituted or unsubstitutedγ-butyrolactone present in the mixture from the first hydrogenation stepis reacted at a temperature lower than the temperature in the firsthydrogenation step to give substituted or unsubstituted tetrahydrofuran.

For the purposes of the present invention the term C₄-dicarboxylic acidsand their derivatives refers to maleic acid and succinic acid which maybear one or more C₁-C₆-alkyl substituents and also the anhydrides andesters of these unsubstituted or alkyl-substituted acids. An example ofsuch an acid is citraconic acid. Preference is given to using MA.

It has surprisingly been found that the use of the Cr-free hydrogenationcatalysts arranged in series and adherence to particular reactionconditions enable the product ratio of GBL:THF to be varied within widelimits.

In this process, a Cr-free catalyst based on copper oxide is used inboth reaction zones. The copper oxide is present in amounts of from 5 to100% by weight. The catalyst can further comprise one or more metals orcompounds thereof, for example oxides, selected from the groupconsisting of Al, Si, Zn, La, Ce, the elements of groups IIIA to VIIIAand of groups IA and IIA in amounts of from 0 to 95% by weight. Thecatalysts used in the two reaction zones may be identical or can havedifferent compositions. Which of these two embodiments is chosendepends, for example, on the desired product composition.

For the purposes of the present invention, the group of the PeriodicTable of the Elements is designated in accordance with the old IUPACnomenclature.

The catalysts used in the first hydrogenation step preferably comprisefrom 5 to 100% by weight of CuO, from 0 to 80% by weight of ZnO and from0 to 95% by weight of Al₂O₃, in particular from 20 to 80% by weight ofCuO, from 10 to 40% by weight of ZnO and from 5 to 60% by weight ofAl₂O₃. The catalysts employed in the second hydrogenation steppreferably comprise from 5 to 80% by weight of CuO, from 0 to 80% byweight of ZnO and from 0 to 60% by weight of Al₂O₃, in particular from20 to 60% by weight of CuO, from 0 to 60% by weight of ZnO and from 10to 50% by weight of Al₂O₃.

The respective metals may also be present in elemental form in the mixedoxide used according to the invention. These are formed, in particular,under a reducing hydrogen atmosphere. In general, the catalyst issubjected to activation, in general pretreatment with hydrogen, beforeuse in the reaction. This produces the active catalyst species. It isgenerally achieved by partial reduction of the oxides present in thecatalyst mixture to the elemental metal which is active in the catalyticreaction occurring in the process of the present invention.

The first hydrogenation step is preferably carried out at ≧200° C., inparticular from 230 to 300° C. The primary product of this firsthydrogenation is substituted or unsubstituted succinic anhydride (SA)which is then hydrogenated further to substituted or unsubstituted GBL.For this reason, deactivation of the catalyst due to coating withrelative nonvolatile SA is observed when the temperature is too low,i.e. the hydrogenation is carried out at below 200° C.

The reaction is carried out so that the starting mixture is reacted withhydrogen in the first reaction zone to form a product mixture comprisingpredominantly substituted or unsubstituted GBL. For this purpose, thestarting mixture is vaporized and passed through the reactor togetherwith a hydrogen-containing gas stream. Here, the proportion of hydrogenin the gas stream is preferably high. It is possible for other gaseouscomponents such as water vapor, hydrocarbons such as methane, ethane orn-butane or carbon monoxide to be present. The reaction conditions(temperature, pressure, GHSV, MA concentration at the inlet) and thecatalyst are chosen so that the GBL yield is maximized while theformation of SA or overhydrogenation products occurs to a minor extent.GBL yields of at least about 30% are sought. GBL yields are preferablyat least about 50%, in particular at least about 70%. Excessively hightemperatures promote the formation of undesirable by-products.

The concentration of the starting material is from 0.1 to 5% by volume,preferably from 0.2 to 3% by volume. At significantly higherconcentrations, the starting material condenses in the reactor and coatsthe catalyst with a liquid film. This is observed in particular in thecase of MA. Significantly higher concentrations would reduce thespace-time yield and make the process unnecessarily expensive. The GHSV(gas hourly space velocity=volume flow of the reaction gas at STPdivided by the bed volume of catalyst) is set so that the startingmaterials and the SA are reacted completely; it is preferably set tofrom 100 to 10000 h⁻¹. The pressure is from 0.5 to 100 bar, preferablyfrom 1 to 50 bar, in particular <20 bar. Although higher pressures aidthe reaction of the starting materials, they also increase the costs ofthe process. Suitable reactors are tube reactors, shell-and-tubereactors in which the catalyst is present as a fixed bed andfluidized-bed reactors.

In the second reaction zone, the temperature is set to a value belowthat in the previous hydrogenation step. The second hydrogenation ispreferably carried out at ≦280° C., preferably from 150 to 240° C.Excessively high temperatures lead to formation of by-products byoverhydrogenation and thus to a reduction in the yield. The gas leavingthe first reaction zone is generally passed to the second zone,preferably without further work-up. Since the temperature of the secondstep is lower than that of the first, the reaction gas should be cooledto the temperature of the second step. Reactors suitable for the secondstep are likewise tube reactors, shell-and-tube reactors orfluidized-bed reactors. The reaction conditions (temperature, pressure,GHSV) and the catalyst in the second reaction zone are chosen so thatGBL is converted into substituted or unsubstituted THF in the desiredselectivity ratio. Temperatures which are too low lead to an unnecessarydrop in space-time yield of the catalyst. For the GHSV and for thepressure, the same ranges given for the first reaction zone apply. Thecomposition of the reaction gas on entering the second stage depends onthe conditions in the first stage. Accordingly, the GBL concentration ispreferably from 0.2 to 2.0% by volume. The product mixture can beseparated by methods known to those skilled in the art; the excesshydrogen can be circulated and reused for the hydrogenation.

In one variant of the invention, both reaction zones are accommodated inone reactor. Suitable reactors are tube reactors, shell-and-tubereactors or combinations of these. One or more heating circuits can beused to set the preferred temperatures in the two reaction zones. Thehydrogenation of MA, maleic acid or its esters to form GBL and THF isassociated with liberation of a large quantity of heat. The reactionenthalpy of the reaction to form GBL is higher than that of thehydrogenation of GBL to THF. In a reactor which is not operatedisothermally, the temperature in the front part of the reactor willtherefore be higher than in the later part. Particular preference istherefore given to generating a temperature profile along thelongitudinal axis of the reactor by means of measures known to thoseskilled in the art so that the preferred temperatures are set in the tworeaction zones of the reactor. The shape of the temperature profiledepends on parameters known to those skilled in the art, for example thevolume-based catalyst activity, the reaction conditions (pressure, GHSVand inlet concentration of the starting materials) and the geometry andthermostating of the reactor.

The process of the present invention makes it possible to achieve(GBL+THF) space yields of ≧98%. The GBL/THF ratio can be varied in arange of from about 90:10 to 0:100.

EXAMPLE

For the first reaction zone, 100 ml of a catalyst composed of 70% byweight of CuO, 25% by weight of ZnO and 5% by weight of Al₂O₃ were mixedwith 100 ml of glass rings of the same size and placed in a tubereactor. The reactor was heated and the reaction gas flowed through itfrom the top downward. MA was pumped as a melt into a vaporizer operatedat 200° C. where it was vaporized in a stream of hydrogen. TheMA/hydrogen mixture having an MA concentration of 1.0% by volume wasthen passed through the reactor. To preheat the gas to the reactiontemperature, a bed of 100 ml of glass rings was introduced above thecatalyst bed.

Before feeding in the MA/hydrogen mixture, the catalyst was subjected topretreatment with hydrogen. For this purpose, the reactor was firstlyflushed with 200 standard 1/h of nitrogen at atmospheric pressure and atthe same time heated over a period of one hour to a temperature in thecatalyst bed of 180° C. The volume flow of nitrogen was then increasedto 950 standard 1/h and an additional 50 standard 1/h of hydrogen wasfed in. A slight temperature increase in the catalyst bed to about 250°C. was observed as a result. After the temperature in the overallcatalyst bed had cooled to 190° C., the volume flow of nitrogen wasgradually reduced to 500 standard 1/h and the hydrogen flow wasincreased to 500 standard 1/h. Finally, the nitrogen flow was switchedoff and the hydrogen flow was increased to 600 standard 1/h.

The reaction was carried out at 5 bar and 240° C. The GHSV was 3000 h⁻¹.

At complete MA conversion, no SA was detected in the gas leaving thereactor. The selectivities to GBL and THF were 88 and 10%, respectively.Overhydrogenation products (mainly butanol and butane) were formed witha selectivity of 2%.

The gases leaving the reactor were then passed to the second reactionzone. This was produced by mixing 100 ml of a catalyst composed of 40%by weight of CuO, 40% by weight of ZnO and 20% by weight of Al₂O₃ with100 ml of glass rings of the same size and placing this mixture in atube reactor. Before carrying out the reaction, the catalyst wasactivated by means of the above-described pretreatment with hydrogen.The reaction gases from the first stage were mixed with a stream ofhydrogen in a vaporizer operated at 150° C. to adjust the temperaturebefore feeding the gases into the second reaction zone. TheGBL/THF/hydrogen gas mixture having a composition of 1.0% by volume ofGBL and 0.1% by volume of THF was then passed through the reactor. Thereaction was carried out at 190° C. and 5 bar. The GHSV was 3000 h⁻¹.

All the GBL present in the reaction gas from the first reaction zone wasreacted completely. The yield of THF based on the GBL fed in was >99%.No by-product formation was observed.

The GBL conversion can be reduced by lowering the temperature, as aresult of which the reaction gas contains more GBL and less THF. In theextreme case, the temperature can be reduced so far that no conversionof GBL occurs.

We claim:
 1. A process for the gas-phase hydrogenation ofC₄-dicarboxylic acids or their derivatives over a catalyst based oncopper oxide to give substituted or unsubstituted γ-butyrolactone and/ortetrahydrofuran, which comprises a first reaction zone in which theC₄-dicarboxylic acid and/or its derivatives is/are reacted to give amixture comprising a substituted or unsubstituted γ-butyrolactone asmain product and a subsequent second reaction zone in which thesubstituted or unsubstituted γ-butyrolactone present in the mixture fromthe first hydrogenation step is reacted at a temperature lower than thetemperature in the first hydrogenation step to give substituted orunsubstituted tetrahydrofuran.
 2. A process as claimed in claim 1,wherein the catalyst comprises from 5 to 100% by weight of copper oxideand from 0 to 95% by weight of one or more metals or their compoundsselected from the group consisting of Al, Si, Zn, La, Ce, the elementsof groups IIIA to VIIIA and groups IA and IIA as active composition. 3.A process as claimed in claim 1, wherein the reaction in the first zoneis carried out at ≧200° C. and the reaction in the second zone iscarried out at ≦280° C.
 4. A process as claimed in claim 3, wherein thereaction in the first zone is carried out at from 230 to 300° C. and thereaction in the second zone is carried out at from 150 to 240° C.
 5. Aprocess as claimed in claim 1, wherein the catalysts used in the tworeaction zones have the same composition.
 6. A process as claimed inclaim 1, wherein the catalysts used in the two reaction zones havedifferent compositions.
 7. A process as claimed in claim 1, wherein thecatalyst used in the first reaction zone comprises from 5 to 100% byweight of CuO, from 0 to 80% by weight of ZnO and from 0 to 95% byweight of Al₂O₃ and the catalyst used in the second reaction zonecomprises from 5 to 80% by weight of CuO, from 0 to 80% by weight of ZnOand from 0 to 60% by weight of Al₂O₃.
 8. A process as claimed in claim7, wherein the catalyst in the first reaction zone comprises from 20 to80% by weight of CuO, from 10 to 40% by weight of ZnO and from 5 to 60%by weight of Al₂O₃, and the catalyst in the second reaction zonecomprises from 20 to 60% by weight of CuO, from 0 to 60% by weight ofZnO and from 10 to 50% by weight of Al₂O₃.
 9. A process as claimed inclaim 1, wherein both reaction zones are accommodated in a reactor whichis not operated isothermally and in which the temperatures are set sothat the temperature in the first reaction zone is higher than that inthe second reaction zone.
 10. A process as claimed in claim 1, whereinthe pressures set in the two reaction zones are, independently of oneanother, from 0.5 to 100 bar.
 11. A process as claimed in claim 10,wherein the pressures are from 1 to 50 bar.
 12. A process as claimed inclaim 10, wherein the pressures are from <20 bar.
 13. A process asclaimed in claim 1, wherein maleic anhydride is used as startingmaterial.
 14. A process as claimed in claim 13, wherein the maleicanhydride concentration in the first reaction step is from 0.1 to 5% byvolume.
 15. A process as claimed in claim 14, wherein the concentrationis from 0.2 to 3% by volume.
 16. A hydrogenation process as claimed inclaim 1, wherein the first and second reactions are carried out,independently of one another, in a tube reactor, a shell-and-tubereactor or a fluidized-bed reactor.
 17. A process as claimed in claim 1,wherein the catalyst comprises from 5 to 100% by weight of copper oxideand from 0 to 95% by weight of one or more metals or their oxidesselected from the group consisting of Al, Si, Zn, La, Ce, the elementsof groups IIIA to VIIIA and groups IA and IIA as active composition. 18.A process as claimed in claim 1, wherein the reaction in the first zoneis carried out at from 230 to 300° C., and the reaction in the secondzone is carried out at from 150 240° C.